Apparatus for separating organic liquid solutes from their solvent mixtures

ABSTRACT

Process and apparatus for extracting an organic liquid from an organic liquid solute/solvent mixture. The mixture is contacted with a fluid extractant which is at a temperature and pressure to render the extractant a solvent for the solute but not for the solvent. The resulting fluid extract of the solute is then depressurized to give a still feed which is distilled to form still overhead vapors and liquid still bottoms. The enthalpy required to effect this distillation is provided by compressing the still overhead vapors to heat them and indirectly to heat the still feed. The process is particularly suitable for separating mixtures which form azeotropes, e.g., oxygenated hydrocarbon/water mixtures. The energy required in this process is much less than that required to separate such mixtures by conventional distillation techniques.

This application is a division of application Ser. No. 079,935, filedSept. 28, 1979 now U.S. Pat. No. 4,349,415.

This invention relates to a process and apparatus for solvent extractionand more particularly to a process and apparatus for extracting largevolumes of liquid organics from solution.

In the commercial processes used for manufacturing many of thehigh-volume, liquid organic compounds such as oxygenated hydrocarbons,it is necessary, usually as a final step, to separate the organiccompounds from aqueous solutions. In many of these mixtures waterconstitutes a major portion of the solution; and in a large number ofthese cases the water and organic liquids form azeotropes. Thus, theseparation of many of these organic compounds from water requiresrelatively large and complex distillation equipment and demands a heavyexpenditure of energy. Likewise, petroleum fuel fractions and lighthydrocarbons must be separated from other organics such as higherboiling hydrocarbons with which they are miscible.

At present, about 3% of the total national energy consumption in theUnited States is used for distillation processes in petroleum refiningand chemical production. It is therefore obvious that if a process andapparatus could be provided which materially decreased the energyrequirements for separating even a portion of such solutes from theirsolutions, the savings in energy would be highly desirable.

It is therefore a primary object of this invention to provide animproved process for extracting liquid organic solutes such as petroleumfuel fractions, straight-run petroleums, light hydrocarbons andaromatics from organic solvents and oxygenated hydrocarbons and the likefrom admixtures with water. It is another object of this invention toprovide a process of the character described which requires less energyinput per unit of organic liquid extracted than is now required in thepresently used distillation processes. An additional requirement is toprovide such a process which makes it possible to employ distillationequipment having fewer stages in smaller and less complex distillationequipment than now used. Yet a further object is provide a process forextracting such liquid organic solutes from their solvents using liquidor supercritical carbon dioxide as an extractant which makes it possibleto take advantage of many of the unique physical properties of thisextractant including favorable diffusion coefficients, low viscosity andlow heat of vaporization. A still further object of this invention is toprovide such a process which uses a fluid extractant, i.e., carbondioxide, which is nonpolluting, nontoxic and relatively inexpensive.

Another primary object of this invention is to provide improvedapparatus for extracting organic liquid solutes from their solutions,the improvement lying in a combination of apparatus components. Anadditional object is to provide apparatus of the character describedwhich makes possible the use of a fluid solvent with resulting savingsin energy requirements.

Other objects of the invention will in part be obvious and will in partbe apparent hereinafter.

The invention accordingly comprises the several steps and the relationof one or more of such steps with respect to each of the others, and theapparatus embodying features of construction, combinations of elementsand arrangement of parts which are adapted to effect such steps, all asexemplified in the following detailed disclosure, and the scope of theinvention will be indicated in the claims.

For a fuller understanding of the nature and objects of the invention,reference should be had to the following detailed description taken inconnection with the accompanying drawings in which

FIG. 1 illustrates the near critical/supercritical regime of carbondioxide and the solubility of naphthalene within this regime;

FIG. 2 is a plot of the relationship between carbon number ofrepresentative organic liquids to be extracted from water and the CO₂-water distribution coefficient for the liquids;

FIG. 3 is a plot of an exemplary vapor recompression cycle for carbondioxide on a fragment of a temperature-entropy diagram for carbondioxide;

FIG. 4 is a detailed flow chart of the method of this invention usingcarbon dioxide as the extractant and an oxygenated hydrocarbon, e.g.,ethanol in water as the solution; and

FIG. 5 is a diagram of the apparatus and system of this invention.

According to one aspect of this invention there is provided a processfor separating an organic liquid from an organic liquid/solvent mixture,comprising the steps of contacting a mixture of an organic liquid soluteand a solvent for said solute with an extractant fluid under conditionsof temperature and pressure to render said extractant fluid a solventfor the organic liquid solute but not for said solvent, thereby forminga fluid extract of the organic liquid solute in the extractant fluid anda raffinate comprising said solvent with minor amounts of the extractantfluid and organic liquid solute; separating the fluid extract and theraffinate; reducing the pressure on the fluid extract to a level to forma two-phase still feed; distilling the still feed to produce a stilloverhead and liquid still bottoms; recompressing the still overhead toprovide recompressed vapor at an elevated temperature; effectingindirect heat exchange between the recompressed vapor and the stillbottoms to provide the thermal energy required in the distilling stepand to form a liquid condensate of the vapor; and recovering a liquidbottoms product comprising the liquid solute.

In a preferred embodiment of this process the fluid extractant isrecovered for recycling. In a further preferred embodiment carbondioxide is used as the fluid extractant.

According to another aspect of this invention there is providedapparatus for separating an organic liquid from an organicliquid/solvent mixture, comprising in combination pressure vessel meansfor effecting contact between a mixture of an organic liquid solute anda solvent for the solute and a pressurized extractant fluid which is asolvent for the organic liquid solute but not for the solvent to producea fluid extract of the organic liquid solute and a raffinate comprisingsolvent with minor amounts of the extractant fluid and the organicliquid solute; distillation vessel means capable of separating a stillfeed into an overhead vapor and liquid bottoms and having associatedtherewith reboiler means including heat exchange means for circulating aheat transfer fluid therethrough in indirect heat exchange relationshipwith the liquid bottoms; first pressure reducing means; first pressureline means incorporating the pressure reducing means arranged to conveythe fluid extract at a reduced pressure as a two-phase still feed fromthe pressure vessel means to the distillation vessel means; vaporcompressor means; second pressure line means incorporating the vaporcompressor means arranged to convey the overhead vapor to the vaporcompressor means and recompressed vapor therefrom to the heat exchangemeans; second pressure reducing means; separator means; third pressureline means incorporating the second pressure reducing means arranged toconvey the still bottoms from the reboiler means to the second pressurereducing means and decompressed still bottoms therefrom to the separatormeans; and means to recover the still bottoms from the separator meansas product organic liquid solute.

The ability of carbon dioxide as a liquid in its near critical state andas a fluid in its supercritical state to serve as an extracting solventhas been known for a number of years. (See for example Francis, A. W.,J. Phys. Chem. 58, 1099 (1954) and Ind. Eng. Chem. 47, 230 (1955).) Nearcritical and supercritical fluids, including carbon dioxide, have beensuggested as solvents for a wide range of materials including variousoils (U.S. Pat. Nos. 1,805,751, 2,130,147, 2,281,865); flavor components(U.S. Pat. No. 3,477,856); caffein in coffee (U.S. Pat. No. 3,843,832);cocoa butter from a cocoa mass (U.S. Pat. No. 3,923,847); fats fromgrains and the like (U.S. Pat. No. 3,939,281); residual hexane fromde-fatted grain (U.S. Pat. No. 3,966,981); and a variety of materialssuch as paraffins, glycerol, oils and fats from a variety ofcompositions (U.S. Pat. No. 3,969,196). A very detailed review of thegeneral field of extraction with supercritical gases is to be found inAngewandte Chemie--International Edition in English, 17: 10, pp 701- 784(October 1978). Of particular interest is the flow sheet of a pilotplant for continuous "destraction" of petroleum top-residues withpropane appearing as FIG. 5 on page 707 of the Angewandte Chemiereference.

Despite the fact that the solvation properties of gases in their nearcritical and supercritical states, and especially of liquid andsupercritical carbon dioxide, have been known, the application of suchproperties has not been made on any commercial scale to the large-volumeorganic liquids; and more importantly it has not been made in a mannerto materially reduce energy requirements below that point at which theadded costs incurred in handling elevated gas pressures are more thanoffset to provide an appreciable overall net savings. The process andapparatus of this invention make the attainment of such net savingspossible.

Many compounds which are gases at ambient temperature and pressure canbe converted to supercritical fluids by subjecting them to conditionssuch that they are at or above their critical pressures andtemperatures. At pressures and/or temperatures somewhat below thecritical points, most of these gases may be liquified to attain what istermed their near-critical state. These gases in their near-criticalliquid or supercritical fluid state become good solvents for manyorganic materials. It is therefore feasible to refer to them as being ina solvent condition, the actual temperature and pressure for any onefluid in its solvent condition being readily determinable for the soluteto be separated and recovered.

Among those gases which may be converted to the solvent-condition fluidstate are hydrocarbons such as methane, ethane, propane, butane,ethylene, and propylene; halogenated hydrocarbons such as thehalomethanes and haloethanes; and inorganics such as carbon dioxide,ammonia, sulfur dioxide, nitrous oxide, hydrogen chloride and hydrogensulfide. Suitable mixtures of these gases may also be used.

Of these gases which may be in the solvent condition, carbon dioxide,ethylene and ethane may be used as illustrative of the temperatures andpressures required. These gases are of particular interest because theyfall within the near-critical and supercritical regimes at essentiallyambient temperature and have critical pressures in the range of 50 to 75atmospheres--pressures which are readily handled by existing equipmentcomponents. The critical temperature and pressure for each of thesegases are well known and, as noted, the solvent condition temperatureand pressure ranges can readily be determined. For example, carbondioxide has a critical temperature of 31° C. and its solvent conditiontemperature may range between about -40° C. and about 150° C. Thecritical pressure of carbon dioxide is 73 atmospheres and its solventcondition pressure may range between about 30 and 150 atmospheres.

Illustrative of the solvent powers of carbon dioxide in the solventcondition is the diagram for the solubility of naphthalene in carbondioxide shown in FIG. 1. It will be seen that within the regime plottedin FIG. 1, the carbon dioxide has solvent properties similar to those ofnormal liquids.

Carbon dioxide in its solvent condition is a preferred fluid solventextractant in the practice of this invention, for it possesses a uniquecombination of properties. In addition to its good solvent propertiesunder the conditions used, it has distinctly favorable diffusioncoefficients compared to normal liquids, a property which gives rise tohigh mass-transfer coefficients. This in turn offers the possibility ofminimizing or even eliminating any significant transport resistance inthe carbon dioxide phase resulting in an increase in the overallextraction rate. It also thereby offers the possibility of decreasingthe size and more effectively optimizing the design of the distillationcolumns used.

A second favorable property of solvent-condition carbon dioxide is itslow viscosity which is about a factor of ten less than that ofconventional liquid solvents. Since viscosity enters into the floodingcharacteristics of an extraction column, high flooding velocities andthus higher flow capacities can be achieved with a concomitant reductionin distillation column diameter.

The high volatility of carbon dioxide relative to many of thelarge-volume organic liquids, e.g., ethanol, methyl ethyl ketone, andthe like which are to be extracted from a water mixture, means that thedistillation column may operate as an evaporator with a short strippingsection using fewer stages. Most important, the vapor rate, and thus theboiler heat requirement is low. Moreover, the heat of vaporization ofthe solvent-condition carbon dioxide is very low--being about one-fifthof that of many normal liquid solvents and about one-thirteenth that ofwater.

Finally, carbon dioxide is inexpensive, nonpolluting and nontoxic,requiring no special equipment or procedures for storage and handlingbeyond normal practice for pressure systems.

The use of solvent-condition fluids according to the process of thisinvention is applicable to the extraction of a wide range of organicliquid solutes from their solutions, whether the solvent to be extractedfrom them is water or another organic liquid, so long as the solvent isrelatively immiscible with the fluid extractant under the conditions oftemperature and pressure employed. Such organic liquid solutes include,but are not limited to, petroleum fuel fractions derived from catalyticcracking and hydrocracking, straight-run petroleum fractions and lighthydrocarbons; aromatics such as styrene and o-xylene; and water-miscibleoxygenated hydrocarbons including the aliphatic alcohols such asethanol, isopropanol and the like; the polyhydric alcohols; as well asacids, aldehydes, esters and ketones.

Since the separation of oxygenated hydrocarbons from water mixtureconstitutes an important commercial process, the extraction of thisclass of solute from an aqueous solution will be used hereinafter asillustrative of the process and apparatus of this invention. Moreparticularly, ethanol is taken as an example of a liquid organic solute.Ethanol is totally miscible with and forms an azeotrope with water whichcontains 89.4 mol % ethanol. The energy consumed in the distillation ofthis mixture is 9008 Btu per pound of alcohol product. The 1976 salesvolume in the United States of synthetic ethanol was 890×10⁶ pounds,indicating that some 8×10¹² Btu were consumed in the separation ofsynthetic ethanol/water mixtures. It becomes obvious from this oneillustration alone that the reduction in the energy required to producesuch organic liquid intermediates as ethanol is highly desirable.

In the practice of this invention it is necessary to chose a solventcondition fluid extractant which exhibits an extractant/waterdistribution coefficient for the organic liquid solute of sufficientmagnitude to ensure that the organic liquid solute will be picked up inthe extractant in preference to the water. Generally a distributioncoefficient of at least 0.1 under the conditions of temperature andpressure used is preferred. These distribution coefficients may readilybe determined either from the literature or by simple experimentation inorder to use the optimum conditions for any given extractant-organicliquid system. For example, it will be seen from FIG. 2, which is a plotof the relationship between distribution coefficient and number ofcarbon atoms in normal aliphatic alcohols and in esters, that thiscoefficient increases rapidly with carbon number. However, even withdistribution coefficients less than one, as in the case of ethylalcohol, the process of this invention can provide material savings inenergy as discussed below.

An important feature of the process of this invention is the use ofsolvent extractant vapor recompression in combination with the use of asolvent-condition fluid extractant. This makes possible the utilizationof the overhead vapor enthalpy as the boiler heat source. In order toaccomplish this, the temperature at which the heat is delivered from thevapor must be raised to provide a ΔT driving force for heat transfer tothe still bottoms in the boiler. This is achieved by vapor compression,so that condensation and enthalpy release will occur at a temperaturehigher than the boiling point of the boiler liquid.

Again using carbon dioxide as exemplary of the solvent-condition fluidextractant, it is possible to show a typical vapor-recompression cycleon the carbon dioxide temperature-entropy diagram of FIG. 3. In thisexample, the solvent-condition carbon dioxide leaving the extractioncolumn is at point A, here taken to be 25° C. and 65 atmospheres whichmeans that the extractant is being used in its near critical liquidstate. Upon expansion into the distillation column, the streamconstituting the still feed drops in pressure at constant enthalpy to 50atmospheres. This is point B which in this example represents about 22%vapor and 78% liquid at 15° C. In the reboiler, enthalpy is added andliquid is vaporized to point C, representing all vapor at the samepressure and temperature. Finally, this vapor, passing overhead from thedistillation column, is then compressed to point D and, in giving upenthalpy in the reboiler, the stream returns from point D to point A.

The steps of the process of this invention are detailed in the flowchart of FIG. 4 and the apparatus is diagrammed in FIG. 5. Referenceshould be had to both of these drawings in the following detaileddescription. Again, carbon dioxide is used for purposes of illustrationas the extractant and ethyl alcohol as the liquid organic solute.

The organic liquid/water mixture feed is pressurized and pumped by pump10 through a suitable pressure line 11 into a pressure vessel 12designed to provide for the contacting of the feed mixture with thesolvent-condition gas extractant introduced into pressure vessel 12through line 13. For convenience of describing this process andapparatus, it will be assumed, for illustrative purposes only, that thefeed mixture is water/ethanol and the fluid extractant is carbondioxide. The extractor 12 may be any suitable pressure vessel designedto provide efficient liquid-liquid contact, such as by countercurrentflow in a packed or sieve-plate tower.

The liquid raffinate, comprised of water, carbon dioxide and a verysmall residual amount of ethanol, is withdrawn from extractor 12 throughline 14 and a pressure-reducing value 15; and the resulting decompressedraffinate is a two-phase mixture of liquid water, with a small amount ofdissolved carbon dioxide as well as the residual ethanol, and carbondioxide vapor. The water phase is withdrawn through line 17 andpressure-reducing valve 18 to become the raffinate discharge. The carbondioxide forming the vapor phase is transferred from separator 16 by line19 to a vapor holding tank 20 for subsequent reconversion to the solventcondition as detailed below.

The liquid carbon dioxide extract containing the dissolved ethanol iswithdrawn from extractor 12 under the same conditions as obtained in theextractor and transferred by pressure line 25 through pressure reducingvalve 26 to the distillation column 27. The reduction of pressure, e.g.,down to 50 atmospheres, experienced by the carbon dioxide extractproduces a still feed, which is part liquid, part vapor, at a lowertemperature, e.g., about 15° C. The distillation column 27 is providedwith sufficient stages to ensure that essentially all of the ethanolcollects in the reboiler 28 along with liquid carbon dioxide forming thestill bottoms.

It will be appreciated that these operational conditions areillustrative and not limiting. For example, the carbon dioxide extractpressure may be reduced to between about 30 and about 80 atmospheresprior to its introduction into distillation column 27; and the resultingstill feed may range between about 0° and 31° C.

In keeping with an important feature of this invention, the heatsupplied to reboiler 28 is provided through out-of-contact or indirectheat exchange with recompressed carbon dioxide vapor drawn from theoverhead of distillation column 27 and sent through line 29, compressor30, and line 31 into heat exchanger coils 32 in reboiler 28. In analternative embodiment, reducing valve 26 may be replaced by a turbine26a, serving as a means for generating energy, the power output of whichmay be used to furnish at least a portion of the power required to drivecompressor 30 to which turbine 26a may be mechanically linked.

In vapor-recompression evaporation or distillation, the elevation inboiling point of the more-volatile component (here the extractant, e.g.,carbon dioxide) caused by the presence of the less-volatile component(here the liquid organic solute) is important. The still overheadleaving the distillation column 27 through line 29 will be at or nearthe boiling point of the more-volatile component; and the liquid (asolution of the solute and extractant) in reboiler 28 will be at ahigher temperature, the magnitude of the difference in temperaturedepending upon the boiling point elevation due to the presence of thesolute.

The still overhead from distillation column 27 is compressedadiabatically in compressor 30 to add the enthalpy which must betransferred to the reboiler liquid to partially vaporize it whilecooling and condensing the compressed vapor as it passes through heatexchanger 32. Thus the mechanism of vapor-recompression distillationrequires that the still overhead must be heated by compression to atemperature high enough above the reboiler liquid temperature to providean economical temperature-difference driving force to effect thenecessary heat transfer within reboiler 28. Therefore it follows thatthe greater the boiling-point elevation due to the presence of thesolute, e.g., ethanol, the greater is the compression required and thegreater is the excess enthalpy that must be added by the compressor toprovide an economical temperature-difference driving force for heattransfer. The magnitude of this excess can in some cases cause vaporrecompression distillation to be uneconomical.

Since the boiling-point elevation for solutions of ethanol and carbondioxide have not been found in the literature, a first approach involvedthe calculation of this parameter using known principles for colligativeproperties and assuming the applicability of Raoult's Law, a commontechnique for predicting vapor-liquid equilibrium data. Assuming asolution of 50% ethanol in carbon dioxide at 50 atmospheres, thecalculated value for boiling point elevation is approximately 50° C.,i.e., the temperature to which the reboiler liquid must be heatedthrough indirect heat exchange with compressed carbon dioxide in heatexchanger 32 would have to be about 50° C. above the normal boilingpoint of carbon dioxide at 50 atmospheres pressure. However, the actualmeasured value of the boiling point elevation under these conditions isabout 3° C. This great discrepancy between calculated and actual valuesfor boiling point elevation may be attributed to the fact that carbondioxide under the conditions employed does not obey Raoult's Law.

Thus it has been found that there exists an unexpectedly favorable lowvalue for the boiling-point elevation in such carbon dioxide solutionsas employed in the process of this invention. It will, of course beappreciated that such a low boiling-point elevation requires only amoderate increase in still overhead pressure. This means that acomparatively small amount of energy is required to compress the stilloverhead and hence to separate the solute from the liquid carbon dioxideextract. This, in turn, in part, gives rise to the low-energycharacteristics associated with the process of this invention.

Following the example which is used to described FIGS. 4 and 5, thestill overhead vapor sent to the compressor is under essentially thesame conditions, 50 atmospheres and 15° C., which prevails indistillation column 27; while the compressed and heated vapor introducedinto heat exchanger 32 is at 65 atmospheres (essentially the extractionpressure) and 36° C. As will be described below, a portion of thecompressed and heated vapor from compressor 30 may be used to heat theexpanded still bottoms from reboiler 28.

Transfer of heat to the liquid in reboiler 28, through heat exchangewith the compressed and heated vapors, results in the boiling off ofadditional carbon dioxide. Because of its very low heat of vaporization,the heat supplied from the recompressed vapor is sufficient to boil offthe carbon dioxide, a fact which results in the material reduction inenergy requirements compared, for example, with the heat required in thedistillation of a liquid organic/water mixture.

The warmed still bottoms are discharged from reboiler 28 through line 35and pressure-reducing valve 36 from which they emerge at a pressure,e.g., 10 atmospheres, intermediate between the still pressure andatmospheric, and at a low temperature, e.g., -40° C. The decompressedcooled still bottoms are then brought back up to a temperature, e.g., toabout 10° C., intermediate between that which they were discharged fromvalve 36 into line 37 and ambient temperature. This heating isaccomplished within heat exchanger 38 using the compressed vapor slipstram drawn off line 31 through line 39 as a heat source. Because it isdesirable to have the two streams of carbon dioxide condensate leavingheat exchanger 28 through line 40 and leaving heat exchanger 38 throughline 41 at or near the extraction temperature, e.g., 28° C., it may benecessary to include a refrigeration means 42 in line 39 to removeenthalpy from the carbon dioxide before recycling it to the extractor.

The still bottoms at the intermediate pressure and temperature arecarried by line 37 into a separator 45 from which the product vaporflash, consisting of carbon dioxide with only very small residualamounts of water and ethanol, is taken by line 46 to vapor holding tank20 to be mixed with raffinate vapor flash. The liquid product ethanol iswithdrawn from separator 45 through line 47, let down to atmosphericpressure in valve 48, and then conveyed as liquid via line 49 to astripping tower 50 from which residual carbon dioxide gas is dischargedthrough line 51 and product ethanol is withdrawn though line 52.

The combined carbon dioxide vapor in holding tank 20 must be convertedto a solvent condition--in this example it must be compressed from 10 to65 atmospheres and delivered to extractor 12 at 28° C. The vapor istherefore taken through line 55 to compressor 56 which is preferably atwo-stage compressor with intercooling. The heat of compression issubsequently removed from the compressed carbon dioxide in one or moreaftercoolers 57 and 58 prior to being carried by line 59 into condensatereturn line 40 which becomes extractant feed line 13. The necessarymake-up solvent-condition carbon dioxide is brought into feed line 13through a pump 60.

It will be apparent from the above description of the invention, asillustrated in FIGS. 4 and 5, that it is possible to carry out theprocess using a wide range of operational parameters so long as certainconditions are met. The fluid used for extracting the organic liquidmust be at a pressure and temperature which make it a solvent for theorganic liquid to be extracted. In selecting an appropriate solventfluid it is preferable that the extractant fluid/water distributioncoefficient of the organic liquid be at least 0.1 for the conditionsused. The choice of conditions used to place the fluid extractant ineither a near critical liquid state or in the supercritical fluid statewill depend upon the physical properties of the gas; upon the solubilitywithin these regimes of the organic liquid solute being extracted; andupon the solubility of the extractant fluid, e.g., carbon dioxide in thesolvent, e.g., water, being removed. Thus the extract withdrawn fromextractor 12 may be a liquid or a supercritical fluid or a combinationof these, the term fluid being used to encompass any one of these forms.Generally it is preferable to choose those pressures and temperaturesapproaching the lower limits of the feasible working ranges because ofthe economics involved, both with respect to original capitalexpenditures and to operating costs.

Because it is necessary to maintain a two-phase system in thedistillation column 27, the pressure of the extract must be reducedbelow the critical pressure of the gas extractant/organic liquid mixtureprior to its introduction into the column. It is, however, desirable tomaintain the pressure differential between the extractor 12 anddistillation column 27 at a relatively low value to minimize the amountof energy required by the system. Such energy is primarily in the formof the compressor work required to return the extractant gas to thepressure used in the extractor.

The temperature of the still feed at its point of introduction in thedistillation column will, of course, be determined by the pressure dropexperienced by the extract in the pressure-reducing valve 26; while thetemperature of the still bottoms must be maintained at the boiling pointof the liquid. Although the still can be operated over a temperaturerange extending from just below the critical temperature of the stillfeed to just above the freezing point of the still bottoms, it ispreferable to operate it as near to ambient temperature as the othernamed operational parameters permit.

The boiling point of the still bottoms, in turn, provides for thedetermination of an optimum temperature or temperature range for thecompressed vapor into the heat exchanger 32 in boiler 28, which, inturn, provides for the determination of the optimum degree ofcompression of the still overhead by compressor 30. It is within theskill in the art to balance this degree of compression and theconcomitant increase in temperature with the design and complexity ofthe heat exchange means within the reboiler.

The temperature of the compressed vapor entering heat exchanger 32, mustof course, be higher than the boiling point of the still bottoms inorder to provide the necessary ΔT heat exchange driving force. It ispreferable that this ΔT be of sufficient magnitude to make it possibleto use efficient but relatively uncomplicated heat exchange means.Essentially all of the heat exchange should take place as the vaporcondenses in the boiler to establish the most thermally efficientsystem.

Finally, the intermediate pressures chosen for separators 16 and 45 willbe those which achieve an optimum balance between the recovery of asmuch of the extractant fluid as possible and the requirement for aslittle work of compression as need be used.

In the conventional distillation of azeotrope-forming mixtures, theresulting product solute may require additional azeotropic distillationin those cases in which the product is leaner in solute than theazeotropic composition. In the process of this invention, however, thefluid solvent and process conditions may be chosen to provide a productsolute which is richer in solute than the corresponding azeotropecomposition, thereby making it possible to eliminate the more difficultand energy-consuming azeotropic distillation step and to substituteconventional distillation for it. Therefore, in some cases wheresufficient solvent remains in the organic liquid solute product, it maybe desirable to subject the product liquid withdrawn through line 52 toa final distillation step in conventional distillation apparatus 53.Such an optional final distillation step will, of course, require farless energy than would be required to effect the separation of theliquid organic solute and solvent solely by conventional distillationfollowed by any necessary azeotropic distillation.

It is also within the scope of this invention to subject the liquidbottoms product discharged from reboiler 28 to a second extraction usingessentially the same process and apparatus as that previously described.Thus as indicated in dotted lines in FIGS. 4 and 5, the pressurizedstill bottoms discharged through line 35 may be taken by way of line 54and pump 55 into a second extractor 12a into which extractant isintroduced through line 13a and carbon dioxide extract is withdrawnthrough line 25a. Since the still bottoms withdrawn from reboiler 28will be at a pressure and a temperature which are somewhat below thepressure and temperature at which extractor 12a operates, somecompression of these still bottoms will be required. It may also benecessary to adjust the temperature of the resulting compressedextractor feed through suitable heat exchange means (not shown).Finally, inasmuch as the still bottoms providing the feed for extractor12a contain some carbon dioxide, the amount of extractant brought intothat extractor is adjusted to take this into account.

The individual apparatus components are either presently available orcan be readily designed and constructed using available informationconcerning materials and performance of related available components. Inthe case of some of the components it may be found desirable to usespecific embodiments or modifications of known equipment to achieve anoptimum design balance in the overall system. Thus, for example, it maybe desirable to use a pulsed extraction column to ensure that the smalldroplets of water making up the discontinuous phase are efficientlysuspended throughout the extractor liquid during contacting andextracting.

Since essentially all of the apparatus components--vessels, lines,valves, heat exchangers, separators, distillation columns andreboiler--must be operated at pressures above ambient, it is desirableto choose as the fluid extractant, a gas having a relatively lowcritical pressure, i.e., below about 100 atmospheres. Likewise thosegases, the critical temperatures of which are relatively low andpreferably not far from ambient temperatures are preferred.

Through the use of the process and apparatus of this invention it ispossible to materially reduce the energy requirements for separatingorganic liquids from their water mixtures. Inasmuch as many such organicliquids are produced in very large volumes, the realization of areduction in even a portion of the energy now required would be highlydesirable.

It will thus be seen that the objects set forth above, among those madeapparent from the preceding description, are efficiently attained and,since certain changes may be made in carrying out the above process andin the constructions set forth without departing from the scope of theinvention, it is intended that all matter contained in the abovedescription or shown in the accompanying drawings shall be interpretedas illustrative and not in a limiting sense.

We claim:
 1. An apparatus for separating an organic liquid from anorganic liquid/solvent mixture, comprising in combination(a) pressurevessel means for effecting contact between a mixture of an organicliquid solute and a solvent for said solute and a pressurized extractantfluid which is a solvent for said organic liquid solute butsubstantially less for said solvent to produce a fluid extract of saidorganic liquid in said extractant fluid and a raffinate comprising saidsolvent with minor amounts of said extractant fluid and said organicliquid solute, said extractant fluid being a gas at ordinary ambientconditions of temperature and pressure; (b) distillation vessel meanscapable of separating a still feed into an overhead vapor and liquidbottoms and having associated therewith reboiler means including heatexchange means for circulating a heat transfer fluid therethrough inindirect heat exchange relationship with said liquid bottoms; (c) firstpressure line means arranged for conveying said fluid extract as a stillfeed from said pressure vessel means to said distillation vessel means;(d) vapor compressor means; (e) second pressure line means incorporatingsaid vapor compressor means and arranged for conveying said overheadvapor to said vapor compressor means and recompressed vapor therefrom tosaid heat exchange means; (f) still bottom pressure reducing means; (g)product separator means; (h) third pressure line means incorporatingsaid still bottom pressure reducing means and arranged for conveyingsaid still bottoms from said reboiler means to said still bottompressure reducing means and decompressed still bottoms therefrom to saidproduct separator means; (i) means for recovering said still bottomsfrom said product separator means as product organic liquid solute; (j)raffinate separating means for providing a raffinate vapor flash; (k)vapor flash collection means; and (l) means disposed for conveying saidraffinate vapor flash and solvent vapor flash from said means forrecovering said still bottom means for recirculation as said solvent. 2.An apparatus in accordance with claim 1 including fluid extract pressurereducing means associated with said first pressure line means.
 3. Anapparatus in accordance with claim 2 wherein said fluid extract pressurereducing means comprises energy generating means.
 4. An apparatus inaccordance with claim 3 wherein said energy generating means aremechanically linked to said vapor compressor means to provide powerthereto.
 5. An apparatus in accordance with claim 2 wherein saidraffinate separating means comprises(1) raffinate pressure reducingmeans for reducing the pressure of said raffinate to a levelintermediate between the pressure in said pressure vessel means andambient pressure so as to produce said raffinate vapor flash and araffinate liquid; (2) raffinate separator means for separating saidraffinate vapor flash from said raffinate liquid; (3) fourth pressureline means incorporating said raffinate pressure-reducing means andproviding fluid communication between said pressure vessel means andsaid raffinate separator means; and (4) fifth pressure line meansarranged for conveying product vapor flash from said raffinate separatormeans to said vapor flash collection means; and wherein said apparatusfurther includes(m) sixth pressure line means arranged for conveyingfluid condensate from said heat exchanger to said pressure vessel means;(n) means for conveying solvent vapor flash from said vapor flashcollection means to said means for recovering said still bottoms; (o)solvent vapor flash compressor means; and (p) seventh pressure linemeans incorporating said solvent vapor flash compressor means andarranged for conveying vapor from said vapor flash collection means tosaid solvent vapor flash compressor means and pressurized extractantfluid therefrom into said sixth pressure line.
 6. An apparatus inaccordance with claim 5 including eighth pressure line means connectingsaid second and third pressure line means and having supplemental heatexchange means for effecting heat exchange between a slip stream of saidrecompressed vapor and said still bottoms.
 7. An apparatus in accordancewith claim 6 including means associated with said eighth pressure linemeans for adjusting the enthalpy in said slip stream prior to its entryinto said supplemental heat exchange means.
 8. An apparatus inaccordance with claim 5 including heat exchange means for adjusting thetemperature of said pressurized extractant fluid in said seventhpressure line.
 9. An apparaus in accordance with claim 5 including meansfor introducing makeup pressurized fluid extractant into said sixthpressure line.
 10. An apparatus in accordance with claim 5 wherein saidproduct separator means includes stripper means, product pressurereducing means, conduit means connecting said separator means with saidstripper means and incorporating said product pressure reducing means,and means for withdrawing finally stripped product organic liquid fromsaid stripper means.
 11. An apparatus in accordance with claim 10including distillation means and means to convey said finally strippedproduct organic liquid from said stripper means to said distillationmeans.
 12. An apparatus in accordance with claim 5 including secondraffinate pressure reducing means for reducing the pressure of saidraffinate liquid from said raffinate separator means to atmospheric. 13.An apparatus in accordance with claim 2 including means for distillingsaid product organic liquid solute so as to remove residual solvent.